Multiple riser reactor

ABSTRACT

The present invention is directed to a hydrocarbon conversion apparatus. The apparatus comprises the following: a plurality of riser reactors, each of the riser reactors having a first end into which a catalyst can be fed and a second end through which the catalyst can exit the riser reactor; a separation zone into which the second ends of the riser reactors extend, the separation zone being provided to separate the catalyst from products of a reaction conducted in the hydrocarbon conversion apparatus; and at least one catalyst return in fluid communication with the separation zone and the first ends of the riser reactors, the catalyst return being provided to transfer the catalyst from the separation zone to the first ends of the riser reactors.

FIELD OF THE INVENTION

The present invention relates to a reactor useful in hydrocarbonconversion processes and particularly in oxygenate to olefin conversionreactions.

BACKGROUND OF THE INVENTION

When converting a feedstock containing a hydrocarbon to a product in anindustrial reactor, it is desirable to maximize the production of adesired product or products, and to control, typically to minimize, theproduction of by-products. One type of reactor useful for conductinghydrocarbon conversion reactions is a fluidized bed reactor, whereinsolid catalyst particles are suspended in a fluidized state duringcontact with the feedstock and other vapor materials. These types ofreactors usually have a cylindrical reactor geometry. One method forreducing the production of by-products in a fluidized bed reactorinvolves operating in a hydrodynamic flow regime such that thesuperficial gas velocity obtains a velocity high enough to cause the netflow of catalyst in the reactor to flow in the same direction as theflow of the feedstock and other vapors, i.e., the feedstock and othervapors essentially carry the catalyst particles along with them. Theseflow regimes are known to those skilled in the art as the fast-fluidizedbed and riser regimes, more generally as the transport regime, and arepreferred in reaction systems in which a more plug flow reactor type isdesired.

In general, for a given reactor cross sectional area (which in acylindrical reactor geometry is proportional to the diameter, and moregenerally to a characteristic width), the catalyst concentration in afluidized bed reactor decreases with increasing gas superficialvelocity. Higher gas superficial velocities generally require tallerreactor heights to allow a given amount of feedstock to contact arequired amount of catalyst. These higher gas superficial velocitiesnecessitate a higher aspect ratio (the ratio of a reactor height to itsdiameter or characteristic width) of the reactor. Further, in many casesit is desired to make a fluidized reactor with a very largecross-sectional area to enable very large throughputs of feedstock in asingle reactor facility. However, increasing fluid bed diameter,particularly in the transport regime, also necessitates increasedreactor height. This increased height is required because a certainminimum reactor height, in terms of a minimum aspect ratio, is requiredto achieve a fully developed flow pattern which approximates plug flowreactor behavior. At the exit and, particularly, at the entrance of atransport regime fluidized bed reactor, unsteady state momentum effectsdominate hydrodynamic behavior (e.g., the energy required for thefeedstock vapors to pick up and accelerate the solid catalyst againstthe force of gravity) in a manner not conducive to obtaining approximateplug flow behavior. Not until these momentum effects are dampened out byprogressing along the reactor height will a well behaved, approximatelyplug flow fluid/solid flow pattern emerge. Finally, should the use oflower activity catalysts be desired in the transport regime, aspectratios must also increase to provide desired higher feedstockconversion.

Unfortunately, high aspect ratio transport fluid bed reactors aredifficult and expensive to construct and maintain. They are expensivebecause they must have at the top a very large, heavy separation vessel,often filled with heavy equipment, to capture and manage the flowingcatalyst and reactor product. As the reactor increases in height (aspectratio), more expensive support structures may be required. In certainareas of the world where inclement, especially windy weather occursroutinely, even more structural support is required, and certain aspectratios are not economic. Multiple, complete and independent reactorsystems with independent separation vessels are required. With thesemultiple, complete and independent reactor systems come attendantmultiplication of costs.

Thus, a need exists in the art for a reactor which can provide thedesired aspect ratio without necessitating an unwieldy height, forcing awidth in which the desired, fully developed flow regime may never beobtained, or without resulting to multiple, independent reactor systems.

SUMMARY OF THE INVENTION

The present invention provides a solution to the currently existing needin the art by providing a hydrocarbon conversion apparatus whichcomprises a plurality of riser reactors. By providing a plurality ofriser reactors, the width or diameter of the feedstock conversionreactor can be reduced, and thus a desired aspect ratio can bemaintained with its attendant closer approach to a desired, fullydeveloped flow regime, at a reduced and more manageable reactor-height.Further, the invention provides the proper aspect ratio for a givenriser reactor without the need for multiple, independent reactorsystems.

One aspect of the present invention is directed to a hydrocarbonconversion apparatus. The hydrocarbon conversion apparatus comprises thefollowing: a plurality of riser reactors, each of the riser reactorshaving a first end into which a catalyst can be fed and a second endthrough which the catalyst can exit the riser reactor; a catalystretention zone provided to contain catalyst which can be fed to theriser reactors; a separation zone into which the second ends of theriser reactors extend, the separation zone being provided to separatethe catalyst from products of a reaction conducted in the hydrocarbonconversion apparatus; a catalyst return in fluid communication with theseparation zone and the catalyst retention zone; and a feed distributorincluding at least one feed head positioned adjacent to each of thefirst ends of the riser reactors.

Another aspect of the present invention is directed to a hydrocarbonconversion apparatus. The apparatus comprises the following: a pluralityof riser reactors, each of the riser reactors having a first end intowhich a catalyst can be fed and a second end through which the catalystcan exit the riser reactor; a separation zone into which the second endsof the riser reactors extend, the separation zone being provided toseparate the catalyst from products of a reaction conducted in thehydrocarbon conversion apparatus; and at least one catalyst return influid communication with the separation zone and the first ends of theriser reactors, the catalyst return being provided to transfer thecatalyst from the separation zone to the first ends of the riserreactors.

Yet another aspect of the present invention is directed to a hydrocarbonconversion process. The process comprises the following steps: (a)contacting a fluidizable catalyst with a fluidizing fluid to fluidizethe fluidizable catalyst; (b) feeding the catalyst and a feed to aplurality of riser reactors, the plurality of riser reactors being partof a single hydrocarbon conversion apparatus; (c) reacting the feed withthe catalyst in the plurality of riser reactors, the reaction of thefeed and the catalyst producing a product; (d) separating the catalystfrom the product in a separation zone, the separation zone being influid communication with the plurality of riser reactors; (e) returningthe catalyst from the separation zone to the plurality of riserreactors; and (f) repeating steps (a) to (e).

These and other advantages of the present invention shall becomeapparent from the following detailed description of the invention, andthe appended drawings and claims.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 presents a partial cross sectional view of a hydrocarbonconversion apparatus of the present invention.

FIG. 2 presents a partial cross sectional view of another embodiment ofthe hydrocarbon conversion apparatus of the present invention.

FIG. 3 presents a partial cross sectional view of yet another embodimentof the hydrocarbon conversion apparatus of the present invention.

FIG. 4 presents a partial cross sectional view of still anotherembodiment of the hydrocarbon conversion apparatus of the presentinvention.

FIG. 5 presents cross sectional views of representative arrangements andconfigurations of the riser reactors and catalyst returns.

DETAILED DESCRIPTION OF THE INVENTION

FIG. 1 presents a partial cross sectional view of a hydrocarbonconversion apparatus 10 of the present invention. The apparatus 10comprises a shell 12, a plurality of riser reactors 20, a feeddistributor 30, and a catalyst return 50.

With continuing reference to FIG. 1, the shell 12 forms a separationzone 14 in which a product of the hydrocarbon conversion reaction isseparated from the catalyst which catalyzes-the-hydrocarbon conversionreaction. Shell 12 includes a first end 16 and a second end 18. Theseparation zone 14 may additionally contain one or more separationdevices, not shown, which are used to separate the products from thecatalyst. Useful separation devices are discussed below in associationwith the discussion of other embodiments of the present invention.Further, the separation devices may be positioned externally to theseparation zone 14, i.e., outside of the shell 12 of the hydrocarbonconversion apparatus 10, or a combination of externally and internallypositioned separation devices.

Riser reactors 20 extend into shell 12 and into the separation zone 14.By extending the riser reactors 20 into shell 12 and the separation zone14, the height required to obtain the desired aspect ratio of a givenriser reactor 20 is concurrent with at least a portion of the heightrequired for the shell 12, separation zone 14, and other associatedspaces, reducing the total height of the hydrocarbon conversion reactor10 of the present invention. Each riser reactor 20 includes a first end22 into which the catalyst and feed are fed to conduct the hydrocarbonconversion reaction. Each riser reactor 20 further includes a second end24 through which the catalyst, products and unreacted feed, if any, exitthe riser reactor 20. The first end 22 of each riser reactor 20terminates in a mouth 26 through which the catalyst and feed are fedinto the riser reactor 20. The number of riser reactors 20 employed inthe hydrocarbon conversion apparatus 10 varies depending on thehydrocarbon conversion process to be conducted in the apparatus 10. Theapparatus 10 can contain two, three, four, five, six and even more thansix riser reactors 20.

The size of the riser reactors 20 depends on parameters such as thesuperficial gas velocity, solids hydrodynamics, pressure, and productioncapacity of the desired hydrocarbon conversion process. In the presentinvention, each riser reactor 20 desirably has a height from 10 metersto 70 meters and a width (or diameter) of one meter to three meters. Allof the riser reactors 20 have a similar height from their first ends 22to their second-ends 24. Desirably, the heights of the riser reactors 20vary by no more than 20% from one riser reactor 20 to another riserreactor 20. More desirably, the heights vary by no more than 10% and,most desirably, the heights vary by no more than 1%.

In the present invention, each of the riser reactors 20 has a similarcross sectional area along its entire height. Desirably, each of theriser reactors 20 has a cross sectional area of no greater than 12 m².More desirably, each of the riser reactors 20 has a cross sectional areaof no greater than 7 m². Most desirably, each of the riser reactors 20has a cross sectional area of no greater than 3.5 m². Desirably, thecross sectional areas of the riser reactors 20 vary by no more than 20%from one riser reactor 20 to another riser reactor 20. More desirably,the cross sectional areas of the riser reactors 20 vary by no more than10% and, most desirably, the cross sectional areas of the riser reactors20 vary by no more than 1%. If one or more riser reactors 20 have both alargest and a smallest cross-sectional area at different points alongthe height of riser reactors 20, desirably the largest cross-sectionalareas of the riser reactors 20 vary by no more than 20% from one riserreactor 20 to another riser reactor 20, and the smallest cross-sectionalareas of the riser reactors 20 vary by no more than 20% from one riserreactor 20 to another riser reactor 20. More desirably, the largestcross sectional area of one riser reactor 20 varies by no more than 10%from the largest cross sectional area of another riser reactor 20 andthe smallest cross sectional area varies by no more than 10% from thesmallest cross sectional area of another riser reactor 20. Mostdesirably, the largest cross sectional area of one riser reactor 20varies by no more than 1% from the largest cross sectional area ofanother riser reactor 20 and the smallest cross sectional area varies byno more than 1% from the smallest cross sectional area of another riserreactor 20.

Desirably, the cross sectional area of each riser reactor 20 varies byno more than 50% along its entire length. More desirably, the crosssectional area of each riser reactor 20 varies by no more than 30% alongits entire height and, most desirably, the cross sectional area of eachriser reactor 20 varies by no more than 10% along its entire height.

To provide a feed to the riser reactors 20, at least one feeddistributor 30 is positioned near the first ends 22 of the riserreactors 20. More than one feed distributor 30 may employed adjacent thefirst ends 22 of the riser reactors 20 to provide feed in variousstates, e.g., one feed distributor 30 may provide feed in a vapor formwhile a second feed distributor 30 may provide feed in a liquid form.Feed distributor 30 includes a body 32 from which a plurality of necks34 extend. Each riser reactor 20 has at least one associated neck 34.Each neck 34 terminates in a head 36. Each head 36 of each neck 34 ispositioned adjacent to the first end 22 of each riser reactor 20.Desirably, each head 36 extends upwardly into each riser reactor 20.More desirably, each head 36 is positioned at or above the mouth 26 atthe first end 22 of each riser reactor 20. Feed distributor 30 mayinclude an optional flow control device, not shown, positioned on feeddistributor 30 to control the amount of feed to each neck 34 or a flowcontrol device may be positioned on each neck 34. The flow controldevice can also be employed to measure flow as well as control it.Further, a nozzle, not shown, may be positioned on each head 36 tofurther control the distribution of the feed to each riser reactor 20.Additionally, each head 36 may be fitted with screening device, notshown, to prevent back flow of catalyst into any of necks 34 of feeddistributor 30 and, subsequently, into body 32 of feed distributor 30.

At least one catalyst return 50 provides fluid communication between theseparation zone of 14 of shell 12 and the riser reactors 20.Particularly, each catalyst return 50 provides fluid communicationbetween the separation zone 14 and the first ends 22 of each riserreactor 20. Each catalyst return 50 has a first end 52 and a second end54. The first end 52 of the catalyst return 50 opens into the second end18 of shell 12 and the second end 54 of catalyst return 50 opensadjacent the riser reactors 20. Each catalyst return 50 is provided totransport catalyst from the separation zone 14 of shell 12 to the firstends 22 of the riser reactors 20. The apparatus 10 may include one, two,three, four, five, six or more catalyst returns 50. Typically, althoughnot necessarily, the number of catalyst returns 50 corresponds to thenumber of riser reactors 20. In the embodiment shown in FIG. 1, thecatalyst returns 50 are external to the riser reactors 20. However, asshown in subsequently described embodiments, the catalyst return 50 maybe contained within a common shell or be positioned internally inrelation to the riser reactors 20 or some combination thereof. Flow ofcatalyst through the catalyst return(s) 50 may optionally be controlledthrough the use of a flow control device 56 positioned on each catalystreturn 50. The flow control device 56 can be any type of flow controldevice currently in use in the art to control catalyst flow throughcatalyst transfer lines. If employed, the flow control device 56 isdesirably a ball valve, a plug valve or a slide valve.

The apparatus 10 further includes a base 60. In the embodiment shown inFIG. 1, the base 60, the catalyst returns 50 and the first ends 22 ofthe riser reactors 20 define a catalyst retention zone 62. The catalystretention zone 62 is provided to retain catalyst which is used tocatalyze the hydrocarbon conversion reaction which is conducted in theapparatus 10. The catalyst return 50 provides fluid communicationbetween the separation zone 14 and the catalyst retention zone 62. To doso, the second ends 54 of the catalyst returns 50 open to the catalystretention zone 62. As one of skill in the art will appreciate, theboundary between the catalyst retention zone 62 and the catalyst return50 is fluid and depends, at least in part, on the level of catalystcontained in the catalyst return 50 and the catalyst retention zone 62.

A fluid distributor 70 is also positioned in or near the base 60 of theapparatus 10. The fluid distributor 70 includes a conduit 72 into whicha fluidizing fluid is fed into catalyst retention zone 62 to fluidize afluidizable catalyst in the catalyst retention zone 62 and the catalystreturns 50. Additional fluid distributors 70, as shown in FIG. 1, mayalso be positioned on each catalyst return 50 to fluidize a fluidizablecatalyst contained in each of the catalyst returns 50.

The hydrocarbon conversion apparatus 10 may also include an outlet 80through which the catalyst can be removed from the apparatus 10. Theoutlet 80 is shown as being positioned on the second end 18 of the shell12 but may be positioned at any position on the apparatus 10. Theapparatus 10 may also include an inlet 82 through which the catalyst maybe placed into the apparatus 10. Although the inlet 82 is shown as beingpositioned on the first end 16 of the shell 12, the inlet 82 may bepositioned at any position on the apparatus 10. A line 84 may beprovided to remove hydrocarbon conversion products from the apparatus10.

As shown in FIG. 1, the hydrocarbon conversion apparatus 10 of thepresent invention may optionally include an associated catalystregeneration apparatus 90. The catalyst regeneration apparatus 90 is influid communication with the hydrocarbon conversion apparatus 10. Thecatalyst regeneration apparatus 90 includes a catalyst regenerator 92,which is in fluid communication with the hydrocarbon conversionapparatus 10, and an optional catalyst stripper 94, which is in fluidcommunication with the catalyst regenerator 92 and which may be in fluidcommunication with the hydrocarbon conversion apparatus 10. A first line96 provides fluid communication between the catalyst stripper 94 and theoutlet 80 on shell 12. A second line 98 provides fluid communicationbetween the catalyst stripper 94 and the catalyst regenerator 92. Athird line 100 provides fluid communication between the catalystregenerator 92 and the inlet 82 on shell 12. A flow control device 102may optionally be positioned on first line 96 to control the flow ofcatalyst between the shell 12 and the catalyst stripper 94. A flowcontrol device 104 may optionally be positioned on second line 98 tocontrol the flow of catalyst between the catalyst stripper 94 and thecatalyst regenerator 92. Finally, a flow control device 106 may bepositioned on third line 100 to control the flow of catalyst between thecatalyst regenerator 92 and the shell 12. The flow control devices 102,104 and 106 can be any types of flow control devices currently in use inthe art to control catalyst flow through catalyst transfer lines. Usefulflow control devices include ball valves, plug valves and slide valves.Although the catalyst stripper 94 is shown on FIG. 1 as being separatefrom the catalyst regenerator 92, one skilled in the art will appreciatethat the catalyst stripper 94 may be integrally formed with the catalystregenerator 92. One skilled in the art will also appreciate that,although FIG. 1 shows third line 100 as returning the catalyst to theseparation zone 14 through line 82, the catalyst may also be returned tothe catalyst return 50, the catalyst retention zone 62 and combinationsof the separation zone 14, the catalyst return 50 and the catalystretention zone 62.

When in operation, the hydrocarbon conversion apparatus 10, as shown inFIG. 1, functions in the following manner. The apparatus 10 is filledwith an appropriate amount of a catalyst suitable to carry out thedesired hydrocarbon conversion reaction. The catalyst should be of atype which is fluidizable. At least a portion of the catalyst iscontained in the catalyst retention zone 62. To fluidize the catalyst inthe catalyst retention zone 62, a fluidizing fluid is fed into the fluiddistributor(s) 70 through inlet 72. The fluidizing fluid is fed into thecatalyst retention zone 62 and the catalyst return(s) 50 of thehydrocarbon conversion apparatus 10. Useful fluidizing fluids include,but are not limited to, inert gasses, nitrogen, steam, carbon dioxide,hydrocarbons, and air. The choice of fluidizing fluid depends upon thetype of conversion reaction being conducted in the hydrocarbonconversion apparatus 10. Desirably, the fluidizing fluid is unreactive(i.e. inert) in the reaction being conducted in the hydrocarbonconversion apparatus 10. In other words, it is desirable that thefluidizing fluid does not play a part in the hydrocarbon conversionprocess being conducted in the hydrocarbon conversion apparatus 10 otherthan to fluidize the fluidizable catalyst.

Once the catalyst has reached an acceptable fluidized state, a feed isfed into the hydrocarbon conversion apparatus 10 through feeddistributor 30. The feed enters the body 32 of feed distributor 30,passes through the necks 34 of feed distributor 30 and exits through theheads 36 of feed distributor 30. The feed is distributed to each of theriser reactors 20 through their first ends 22. Desirably, the feed isprovided in substantially equal streams to each riser reactor 20. By“substantially equal” it is meant that the flow of feed provided to eachriser reactor 20 through the feed distributor 30 varies by no more than25% by volume rate, and varies no more than 25% by mass percent for eachcomponent in the feed, from one riser reactor 20 to another riserreactor 20. More desirably, the flow of feed provided to each riserreactor 20 through the feed distributor 30 varies by no more than 10% byvolume rate, and varies no more than 10% by mass percent for eachcomponent in the feed, from one riser reactor 20 to another riserreactor 20. Most desirably, feed provided to each riser reactor 20through the feed distributor 30 varies by no more than 1% by volumerate, and varies no more than 1% by mass percent for each component inthe feed, from one riser reactor 20 to another riser reactor 20.

A pressure differential created by the velocity of the feed entering thefirst ends 22 of the riser reactors 20 and the pressure of the height offluidizable catalyst in the catalyst return(s) 50 and the catalystretention zone 62 causes catalyst to be aspirated into the first ends 22of the riser reactors 20. The catalyst is transported through the riserreactors 20 under well known principles of eduction in which the kineticenergy of one fluid, in this case the feed, is used to move anotherfluid, in this case the fluidized catalyst. The catalyst and feed travelfrom the first ends 22 to the second ends 24 of the riser reactors 20.As the catalyst and feed travel through the riser reactors 20, thehydrocarbon conversion reaction occurs and a conversion product isproduced.

By designing the hydrocarbon conversion apparatus 10 with thesefeatures, each individual riser reactor 20 operates in a substantiallyidentical manner. With this invention, it is desirable to maintain boththe reactant feed rates and the catalyst feed rates at the same rates toeach of the riser reactors 20. In this way, the conversion of the feedand selectivity to the desired products will be substantially identicaland can run at optimum operational conditions.

The conversion product(s), unreacted feed, if any, and the catalyst exitthe riser reactors 20 through their second ends 24 and enter into theseparation zone 14 of shell 12. In second end 16 of shell 12, theconversion product and unreacted feed, if any, are separated from thecatalyst by a separator, not shown, such as cyclonic separators,filters, screens, impingement devices, plates, cones, other deviceswhich would separate the catalyst from the product of the conversionreaction, and combinations thereof. Desirably, the conversion productand unreacted feed, if any, are separated by a series of cyclonicseparators. Once the catalyst has been separated from the conversionproduct and the unreacted feed, if any, the conversion products andunreacted feed, if any, are removed from the shell 12 through the line84 for further processing such as separation and purification. Thecatalyst, after being separated from the products and unreacted feed,moves from the shell 12 to the catalyst retention zone 62. The catalystexits shell 12 through the first ends 52 of the catalyst returns 50 andmoves through the catalyst returns 50 to the first ends 54 of thecatalyst returns 50 from which the catalyst moves to the catalystretention zone 62. If desired, the flow of catalyst through the catalystreturns 50 can be controlled by the flow control devices 56. If the flowcontrol devices 56 are employed, a height of fluidizable catalyst ismaintained above each flow control device 56 in the catalyst return 50to allow proper function of the flow control device 56.

If necessary or desired, at least a portion of the catalyst can becirculated to the catalyst regeneration apparatus 90, as shown inFIG. 1. Catalyst to be regenerated is removed from the shell 12 thoughthe outlet 80 and transported, if desired, to the catalyst stripper 94through the first line 96. The flow of catalyst between the hydrocarbonconversion apparatus 10 and the catalyst stripper 94 can be controlledby the flow control device 102. In the catalyst stripper 94, thecatalyst is stripped of most of readily removable organic materials(organics). Stripping procedures and conditions for individualhydrocarbon conversion processes are within the skill of a person ofskill in the art. The stripped catalyst is transferred from the catalyststripper 94 to the catalyst regenerator 92 through the second line 98.The flow of catalyst through the second line 98 may optionally becontrolled by the optional flow control device 104. In the catalystregenerator 92, carbonaceous deposits formed on the catalyst during ahydrocarbon conversion reaction are at least partially removed from thecatalyst. The regenerated catalyst is then transferred to the shell 12of the hydrocarbon conversion apparatus 10 through the third line 100.The flow of catalyst through the third line 100 may optionally becontrolled by the flow control device 106. A transport gas is typicallyprovided to the third line 100 to facilitate transfer of the catalystfrom the catalyst regenerator 92 to the hydrocarbon conversion apparatus10. The catalyst is returned to the shell 12 through the inlet 82.

Another embodiment of the hydrocarbon conversion apparatus 110 of thepresent invention is shown in partial cross section in FIG. 2. Theapparatus 110 comprises a shell 120, a plurality of riser reactors 130,a feed distributor 140, and a catalyst return 150.

With continuing reference to FIG. 2, the shell 120 forms a separationzone 122 in which a product of the hydrocarbon conversion reaction isseparated from the catalyst which catalyzes the hydrocarbon conversionreaction. Shell 120 includes a first end 124 and a second end 126. Shell120 defines a quiescent zone 128 from which catalyst can be withdrawnfrom the hydrocarbon conversion apparatus 110.

Riser reactors 130 extend into shell 120 and the separation zone 122.Each riser reactor 130 includes a first end 132 into which the catalystand feed are fed to conduct the hydrocarbon conversion reaction. Eachriser reactor 130 further includes a second end 134 through which thecatalyst, products and unreacted feed, if any, exit the riser reactor130. The first end 132 of each riser reactor 130 terminates in a mouth136 through which the catalyst and feed are fed into the riser reactor130. As described above, the number of riser reactors 130 employed inthe hydrocarbon conversion apparatus 110 varies depending on thehydrocarbon conversion process to be conducted in the apparatus 110. Thenumber and size of the riser reactors 130 is discussed above inconjunction with the description of FIG. 1.

To provide a feed to the riser reactors 130, at least one feeddistributor 140 is positioned near the first ends 132 of the riserreactors 130. More than one feed distributor 140 may employed to providefeed in various states, e.g., one feed distributor 140 may provide feedin a vapor form while a second feed distributor 140 may provide feed ina liquid form. Feed distributor 140 includes a body 142 from which aplurality of necks 144 extend. Each riser reactor 130 has at least oneassociated neck 144. Each head 146 of each neck 144 is positionedadjacent to the first end 132 of each riser reactor 130. Desirably, eachhead 146 extends upwardly into each riser reactor 130. More desirably,each head 146 is positioned at or above the mouth 136 at the first end132 of each riser reactor 130. Feed distributor 140 may include anoptional flow control device, not shown, positioned on feed distributor140 to provide an equal amount of feed to each neck 144 or a flowcontrol device may be positioned on each neck 144. The flow controldevice may also be employed to measure flow as well as control. Further,a nozzle, not shown, may be positioned on each head 146 to furthercontrol the distribution of the feed to each riser reactor 130.Additionally, each head 146 may be fitted with a screening device, notshown, to prevent back flow of catalyst into any of necks 144 of feeddistributor 140 and, subsequently, into body 142 of feed distributor140.

At least one catalyst return 150 provides fluid communication betweenthe separation zone 122 of shell 120 and the riser reactors 130. Eachcatalyst return 150 has a first end 152 and a second end 154. The firstend 152 of the catalyst return 150 opens adjacent the second end 126 ofshell 120 and the second end 154 of catalyst return 150 opens to theriser reactors 130. Each catalyst return 150 is provided to transportcatalyst from the separation zone 122 of shell 120 to the first ends 132of the riser reactors 130. The apparatus 110 may include one, two,three, four, five, six or more catalyst returns 150. Typically, althoughnot necessarily, the number of catalyst returns 150 corresponds to thenumber of riser reactors 130. Flow of catalyst through the catalystreturn(s) 150 may optionally be controlled through the use of flowcontrol devices, not shown, positioned on each catalyst return 150. Theflow control devices can be any type of flow control devices currentlyin use in the art to control catalyst flow through catalyst transferlines. If employed, the flow control device is desirably a ball valve, aplug valve or a slide valve.

The apparatus 110 further includes a base 160. In the embodiment shownin FIG. 2, the base 160, the catalyst returns 150 and the first ends 132of the riser reactors 130 define a catalyst retention zone 162. Thesecond ends 154 of the catalyst returns 150 open to the catalystretention zone 162. The catalyst retention zone 162 is provided toretain catalyst which is used to catalyze the hydrocarbon conversionreaction which is conducted in the apparatus 110. As one of skill in theart will appreciate, the boundary between the catalyst retention zone162 and the catalyst return 150 is fluid and depends, at least in part,on the level of catalyst contained in the catalyst retention zone 162and the catalyst return 150.

A fluid distributor 170 is also positioned in or near the base 160 ofthe apparatus 110. The fluid distributor 170 includes a conduit 172 intowhich a fluidizing fluid is fed into catalyst retention zone 162 tofluidize a fluidizable catalyst contained in the catalyst retention zone162 and the catalyst returns 150. Additional fluid distributors 170, asshown in FIG. 2, may also be positioned on the catalyst return(s) 150 toprovide additional fluidizing fluid in the catalyst return(s) 150.

The hydrocarbon conversion apparatus 110 may also include an outlet 180through which the catalyst can be removed from the apparatus 10. Theoutlet 180 is positioned adjacent the quiescent zone 128 in the secondend 126 of the shell 120. It is desirable for the outlet 180 topositioned such that catalyst can be removed from the shell 120 throughthe quiescent zone 128. The apparatus 110 may also include an inlet 182through which the catalyst may be placed into the apparatus 110.Although the inlet 182 is shown as being positioned on the second end126 of the shell 120, the inlet 182 may be positioned at any position onthe apparatus 110. Lines 184 are provided to remove products andunreacted feed, if any, from the separation zone 122 of the hydrocarbonconversion apparatus 110.

A series of separation devices 186 are shown as being positioned in theseparation zone 122 of shell 120. The separation devices 186 may becyclonic separators, filters, screens, impingement devices, plates,cones or any other devices which would separate the catalyst from theproduct of the conversion reaction.

An impingement device 190 is positioned in the first end 124 of theshell 120. The impingement device 190 is provided to direct catalystleaving the riser reactors 130 away from the second ends 134 of theriser reactors 130 and to limit the amount of catalyst falling back intothe riser reactors 130. Desirably, the impingement device 190 ispositioned opposite the second ends 134 of the riser reactors 130.

A series of supports 192 are also shown in FIG. 2. The supports 192 aremerely shown to be illustrative of one possible means for supporting thehydrocarbon conversion apparatus 110.

As one of skill in the art will appreciate, the hydrocarbon conversionapparatus shown in FIG. 2 functions similarly to that shown in FIG. 1and will not be discussed in detail except to illustrate those featuresnot shown in FIG. 1.

With reference to FIG. 2, catalyst is provided to the catalyst retentionzone 162 and is fluidized in the catalyst retention zone 162 and thecatalyst returns 150 by the fluidizing fluid provided through the fluiddistributor 170. The feed is provided to the riser reactors 130 throughthe feed distributor 140. The amount of feed provided to each of theriser reactors 130 is the same as that described above in conjunctionwith the description of FIG. 1. The catalyst and feed flow upwardlythrough the riser reactors 130, in the same manner as described above inconjunction with the description of then riser reactors 20 in FIG. 1.

With continuing reference to FIG. 2, the catalyst, product and unreactedfeed, if any, exit through the second ends 134 of the riser reactors 130into the separation zone 122 of the shell 120. At least a portion, anddesirably most, of the catalyst contacts the impingement device 190 andis deflected toward the sides of the shell 120. The separators 186separate at least a portion of the catalyst from the product andunreacted feed. The product and unreacted feed are removed from theshell 120 of the hydrocarbon conversion device 10 through the lines 184.The catalyst, which is separated by the separators 186, falls into thequiescent zone 128. The remainder of the catalyst is returned to contactthe feed through the catalyst returns 150.

A portion of the catalyst contained in the quiescent zone 128 can beremoved from the hydrocarbon conversion apparatus 10 and be sent to acatalyst regeneration apparatus via outlet 180, such as catalystregeneration apparatus 90 shown in FIG. 1, or removed from thehydrocarbon conversion apparatus 110 for further processing.Additionally, catalyst in the quiescent zone 128 may spill over into thecatalyst returns 150 and be returned to contact the feed.

Another embodiment of the hydrocarbon conversion apparatus of thepresent invention is shown in FIG. 3. The apparatus 200 comprises ashell 212, a plurality of riser reactors 220, feed distributors 230, anda catalyst return 250.

With continuing reference to FIG. 3, the shell 212 defines a separationzone 214 in which a product of the hydrocarbon conversion reaction isseparated from the catalyst which catalyzes the hydrocarbon conversionreaction. Shell 212 includes a first end 216 and a second end 218.

Riser reactors 220 extend into shell 212 and the separation zone 214.Each riser reactor 220 includes a first end 222 into which the catalystand feed are fed to conduct the hydrocarbon conversion reaction. Eachriser reactor 220 further includes a second end 224 through which thecatalyst, product, and unreacted feed, if any, exit the riser reactor220. The first end 222 of each riser reactor 220 terminates in a mouth226 through which the catalyst and feed are fed into the riser reactor220. The number and dimensions of the riser reactors 220 is discussedabove in conjunction with the description of FIG. 1.

With continuing reference to FIG. 3, to provide a feed to the riserreactors 220, at least one feed distributor 230 is positioned near thefirst ends 222 of the riser reactors 220. More than one feed distributor230 may employed to provide feed in various states, e.g., one feeddistributor 230 may provide feed in a vapor form while a second feeddistributor 230 may provide feed in a liquid form. Each feed distributorincludes a body, not shown, from which at least one neck 232 extends.Each riser reactor 220 has at least one associated neck 232. Each feeddistributor 230 terminates in a head 234. Each head 234 is positionedadjacent to the first end 222 of each riser reactor 220. Desirably, eachhead 234 extends upwardly into each riser reactor 220. More desirably,each head 234 is positioned at or above the mouth 226 of the first end222 of each riser reactor 220. Feed distributor 230 may include anoptional flow control device, not shown, positioned on feed distributor230 to provide an equal amount of feed to each head 234. The flowcontrol device can also be employed to measure flow as well. Further, anozzle, not shown, may be positioned on each head 234 to further controlthe distribution of the feed to each riser reactor 220. Additionally,each head 234 may be fitted with screening device, not shown, to preventback flow of catalyst into any of the feed distributors 230.

In the hydrocarbon conversion apparatus 200 shown in FIG. 3, a singlecatalyst return 250 is positioned centrally in relation to the riserreactors 220. The catalyst return 250 provides fluid communicationbetween the separation zone 214 of the shell 212 and the riser reactors220. The catalyst return 250 has a first end 252 and a second end 254.The first end 252 of the catalyst return 250 opens into the first end214 of shell 212 and the second end 254 of catalyst return 250 opens tothe riser reactors 220. A series of arms 256 are positioned on thesecond end 254 of the catalyst return 250. The arms 256 extend from thecatalyst return 250 to each of the riser reactors 220 and provide fluidcommunication between the catalyst return 250 and the riser reactors220. The number of arms 256 will correspond to the number of riserreactors 220 with each riser reactor 230 having at least onecorresponding arm 256. The catalyst return 250 is provided to transportcatalyst from the separation zone 214 of shell 212 to the first ends 222of the riser reactors 220. Flow of catalyst through the catalyst return250 may optionally be controlled through the use of a flow controldevice 258 positioned on the catalyst return 250 or on each arm 256. Theflow control device(s) 258 can be any type of flow control devicescurrently in use in the art to control catalyst flow through catalysttransfer lines. If employed, the flow control device 258 is desirably aball valve, a plug valve or a slide valve.

In the embodiment shown in FIG. 3, the first end 252 of the catalystreturn 250 and the arms 256 define a catalyst retention zone 262. Thearms 256 of the catalyst return 250 open to the catalyst retention zone262. The catalyst retention zone 262 is provided to retain catalystwhich is used to catalyze the hydrocarbon conversion reaction which isconducted in the apparatus 200. As one of skill in the art willappreciate, the boundary between the catalyst retention zone 262 and thecatalyst return 250 is fluid and depends, at least in part, on the levelof catalyst contained in the catalyst retention zone 262 and the arms256 of the catalyst return 250.

At least one fluid distributor 270 is positioned beneath the catalystretention zone 262. The fluid distributor 270 includes a conduit 272into which a fluidizing fluid is fed to fluidize a fluidizable catalystin the catalyst retention zone 262 and the catalyst return 250.Additional fluid distributors 270, as shown in FIG. 3, may also bepositioned on the catalyst return 250 to further fluidize fluidizablecatalyst contained in the catalyst return 250.

The hydrocarbon conversion apparatus 200 may also include an outlet 280through which the catalyst can be removed from the apparatus 200. Theoutlet 280 is shown as being positioned on the second end 218 of theshell 212 but may be positioned at any position on the apparatus 200.The apparatus 200 may also include an inlet 282 through which thecatalyst may be placed into the apparatus 200. Although the inlet 282 isshown as being positioned on the second end 218 of the shell 212, theinlet 282 may be positioned at any position on the apparatus 200. A line284 may be provided to remove products from the apparatus 200.

A series of separation devices 286 are shown as being positioned in theseparation zone 214 of shell 212. The separation devices 286 may becyclonic separators, filters, screens, impingement devices, plates,cones or any other devices which would separate the catalyst from theproduct of the conversion reaction. The separation devices 286 are shownin FIG. 3 as cyclonic separators 288.

A series of supports 292 are also shown in FIG. 3. The supports 292 aremerely shown to be illustrative of one possible means for supporting thehydrocarbon conversion apparatus 200.

The hydrocarbon conversion apparatus 200 which is shown in FIG. 3functions similarly to that shown in FIGS. 1 and 2. The apparatus 200shown in FIG. 3 functions in the following manner.

The apparatus 200 is filled with an appropriate amount of catalyst whichis retained in the catalyst return 250 and the catalyst retention zone262. The catalyst is fluidized in the catalyst return 250 and thecatalyst retention zone 262 by means of a fluidizing fluid which isprovided to the hydrocarbon conversion apparatus 200 through theconduits 272 of the fluid distributors 270. The flow of catalyst to theriser reactors 220 can be controlled by the flow control devices 258.Feed is provided to the riser reactors 220 through the feed distributors230. The amount of feed provided to the riser reactors 220 is the sameas that discussed above in conjunction with the description of FIG. 1.The feed and the catalyst flow upwardly in the riser reactors 230 by theprinciple of eduction which is also described above.

The catalyst, product and unreacted feed, if any, exit the riserreactors 220 through their second ends 224. The catalyst is separatedfrom the product and any unreacted feed by the separation devices 286.The separated catalyst is fed to the second end 218 of shell 212 whilethe product and any unreacted feed are removed from the apparatusthrough the line 284.

A portion of the catalyst may be removed from the apparatus 200 throughthe outlet 280 and sent to a regeneration apparatus, not shown, orremoved entirely from the apparatus 200. The regenerated catalyst isreturned to the apparatus 200 through the inlet 282.

The separated catalyst enters the first end 252 of the catalyst return250 and is recycled to be reused in the hydrocarbon conversion reaction.The catalyst is returned through the catalyst return 250 to the catalystcontainment area 262 where the catalyst is maintained in a fluidizedstate by the fluidizing fluid provided through the fluid distributors270.

Another embodiment of the hydrocarbon conversion apparatus 300 is shownin FIG. 4. The apparatus 300 comprises a shell 310, a plurality of riserreactors 330, a feed distributor 340 and a fluid distributor 350.

With continued reference to FIG. 4, the shell 310 is formed by a wall312 and is hollow. Shell 310 has a first end 314 and a second end 316.The first end 314 of shell 310 defines a separation zone 318 in whichthe catalyst is separated from the product of the hydrocarbon conversionreaction. The shell 310 further includes a wall extension 320, whichextends upwardly into the first end 314 of shell 310 from the second end316 of shell 310, and a funnel portion 322. The wall extension 320 andthe funnel portion 322 define a quiescent zone 324 in which a portion ofthe catalyst can be retained prior to being removed from the shell 310.

In the embodiment shown in FIG. 4, a plurality of riser reactors 330 arepositioned inside shell 310, as shown in FIG. 4. Each riser reactor 330extends substantially parallel to a longitudinal axis of shell 310 andhas a wall 331. Each riser reactor 330 has a first end 332 and a secondend 334. The first end 332 of each riser reactor 330 is positioned inthe second end 316 of shell 310. The second end 334 of each riserreactor 330 extends into the first end 314 of shell 310. The first end332 of each riser reactor 330 terminates in a mouth 335 through whichthe catalyst and feed are fed into the riser reactor 330. Although thehydrocarbon conversion apparatus 300 is shown in FIG. 4 as containingthree riser reactors 330, apparatus 300 desirably contains two or moreriser reactors 330. The number and size of the riser reactors 330 isdescribed in conjunction with the description of FIG. 1.

With continuing reference to FIG. 4, wall 312 of shell 310 and wall 331of each of the riser reactors 330 define a catalyst retention zone 336.The catalyst retention zone 336 contains the catalyst utilized tocatalyze the hydrocarbon conversion reaction. When the apparatus 300 isin operation, catalyst retention zone 336 contains the catalyst in afluidized state, as will be described in detail below. Wall extension320, wall 312 of the shell 310 and the walls 331 of each of the riserreactors 330 also define a catalyst return 338. The catalyst return 338directs catalyst which has been used in a conversion reaction from theseparation zone 318 in the first end 314 of the shell 310 to thecatalyst retention zone 336. As one of skill in the art will appreciate,the boundary between the catalyst retention zone 336 and the catalystreturn 338 is fluid and depends, at least in part, on the level ofcatalyst contained in the catalyst retention zone 336.

To provide a feed to the riser reactors 330, at least one feeddistributor 340 is positioned near the first ends 332 of the riserreactors 330. More than one feed distributor 340 may employed to providefeed in various states, e.g., one feed distributor 340 may provide feedin a vapor form while a second feed distributor 340 may provide feed ina liquid form. Feed distributor 340 includes a body 342 from which aplurality of necks 344 extend. Each riser reactor 330 has at least oneassociated neck 344. Each neck 344 terminates in a head 346. Each head346 of each neck 344 is positioned adjacent to the first end 332 of eachriser reactor 330. Desirably, each head 346 extends into each riserreactor 330. More desirably, each head 346 is positioned at or above themouth 335 at the first end 332 of each riser reactor 330. Feeddistributor 340 may include an optional flow control device 348positioned on feed distributor 340 to provide an equal amount of feed toeach neck 344 and, if desired, to measure the flow through each neck344. As shown in FIG. 4, the flow control device 348 is a valve 350.Useful types of valves are described above. Further, a nozzle, notshown, may be fitted onto each head 346 to distribute the feed into eachriser reactor 330. Additionally, each head 346 may be fitted withscreening device, not shown, to prevent back flow of catalyst into anyof necks 344 of feed distributor 340 and, subsequently into body 342 offeed distributor 340.

A fluid distributor 350 is also positioned in second end 316 of shell310. The fluid distributor 350 includes a conduit 352 into which afluidizing fluid is fed to fluidize a fluidizable catalyst in thecatalyst retention zone 336 and the catalyst return 338. An optionaldisperser 354 may be positioned between the fluid distributor 350 andthe catalyst retention zone 336 to disperse the fluidizing fluid aboutthe catalyst retention zone 336 and the catalyst return 338. Disperser354 is desirably positioned perpendicular to the longitudinal axis ofshell 310 in the second end 316 of shell 310. Disperser 354 may be ascreen, a grid, a perforated plate or similar device through which thefluidizing fluid is fed to provide even distribution of the fluidizingfluid to the catalyst retention zone 336.

To separate products from the hydrocarbon conversion reaction from thecatalyst, a separator 360 or series of separators 360, may be positionedin first end 314 of shell 310. The separators 360 are shown in FIG. 4 asbeing cyclonic separators 362. Other types of separators 360 such asfilters, screens, impingement devices, plates, cones and other deviceswhich would separate the products from the catalyst may also bepositioned in the first end 314 of shell 310. The number of separators360 depends upon the desired operating efficiency, particle size of thecatalyst, the gas superficial velocity, production capacity, and otherparameters. The products are removed from shell 310 through a line 364or a plurality of lines 364 for further processing such as, for example,separation and purification.

The apparatus 300 may further include an outlet 370 through whichcatalyst may be removed from the shell 310 and an inlet 372 through iswhich catalyst may be placed into shell 310. The positioning of outlet370 and inlet 372 is not critical. However, it is desirable for theoutlet 370 to be positioned such that catalyst can be removed from theshell 310 through the quiescent zone 324.

An impingement device 380 is positioned in the first end 314 of theshell 310. The impingement device 380 is provided to direct catalystleaving the riser reactors 330 away from the second ends 334 of theriser reactors 330 and to limit the amount of catalyst falling back intothe riser reactors 330.

A support 392 is also shown in FIG. 2. The support 392 is merely shownto be illustrative of one possible means for supporting the hydrocarbonconversion apparatus 300.

As shown in FIG. 4, the hydrocarbon conversion apparatus 300 mayoptionally include an associated catalyst regeneration apparatus 90which is in fluid communication with the hydrocarbon conversionapparatus 300. The catalyst regeneration apparatus 90 includes acatalyst regenerator 92, which is in fluid communication with thehydrocarbon conversion apparatus 300 and an optional catalyst stripper94, which is in fluid communication with the catalyst regenerator 92 andwhich may be in fluid communication with the hydrocarbon conversionapparatus 300. A first line 96 provides fluid communication between thecatalyst stripper 94 and shell 310 through outlet 370. A second line 98provides fluid communication between the catalyst stripper 94 and thecatalyst regenerator 92. A third line 100 provides fluid communicationbetween the catalyst regenerator 92 and the inlet 372 on shell 310. Aflow control device 102 may optionally be positioned on first line 96 tocontrol the flow of catalyst between the shell 12 and the catalyststripper 94. A flow control device 104 may optionally be positioned onsecond line 98 to control the flow of catalyst between the catalyststripper 94 and the catalyst regenerator 92. Finally, a flow controldevice 106 may be positioned on third line 100 to control the flow ofcatalyst between the catalyst regenerator 92 and the shell 310. The flowcontrol devices 102, 104 and 106 can be any flow control-devicecurrently in use in the art to control catalyst flow through catalysttransfer lines. Useful flow control devices include ball valves, plugvalves and slide valves. Although the catalyst stripper 94 is shown onFIG. 4 as being separate from the catalyst regenerator 92, one skilledin the art will appreciate that the catalyst stripper 94 may beintegrally formed with the catalyst regenerator 92. One skilled in theart will also appreciate that, although FIG. 4 shows third line 100 asreturning the catalyst to the separation zone 318 through line 372, thecatalyst may also be returned to the catalyst return 338, the catalystretention zone 336 and combinations of the separation zone 318, thecatalyst return 338 and the catalyst retention zone 336.

When in operation, the hydrocarbon conversion apparatus 300, as shown inFIG. 4, functions in the following manner. The catalyst retention zone336 is filled with a catalyst suitable to carry out the desiredhydrocarbon conversion reaction. The catalyst should be of a type whichis fluidizable. To fluidize the catalyst in the catalyst retention zone336 and the catalyst return 338, a fluidizing fluid is fed into thefluid distributor 350 through conduit 352. The fluidizing fluid isdispersed within the shell 310 of the hydrocarbon conversion apparatus300 by the disperser 354. Useful fluidizing fluids include, but are notlimited to, nitrogen, steam, carbon dioxide, hydrocarbons, and air. Thechoice of fluidizing fluid depends upon the type of conversion reactionbeing conducted in the hydrocarbon conversion apparatus 300.

Once the catalyst has reached an acceptable fluidized state, a feed isfed into the hydrocarbon conversion apparatus 300 through feeddistributor 340. The feed enters the body 342 of feed distributor 340,passes through the necks 344 of feed distributor 340 and exits throughthe heads 346 of feed distributor 340. The feed is distributed to eachof the riser reactors 330 through the mouths 335 at the first ends 332of the riser reactors 330.

A pressure differential created by the velocity of the feed entering thefirst ends 332 of the riser reactors 330 and the pressure of the heightof fluidizable catalyst in the catalyst retention zone 336 causescatalyst to be aspirated into the first ends 332 of the riser reactors330. The catalyst is transported through the riser reactors 330 underwell known principles of eduction in which the kinetic energy of onefluid, in this case the feed, is used to move another fluid, in thiscase the fluidized catalyst. The catalyst and feed travel from the firstends 332 to the second ends 334 of the riser reactors 330. As thecatalyst and feed travel through the riser reactors 330, the hydrocarbonconversion reaction occurs and a conversion product is produced.

The conversion product(s), unreacted feed, if any, and the catalyst exitthe riser reactors 330 through their second ends 334 and enter thecatalyst separation zone 318 in the first end 314 of shell 310. In thecatalyst separation zone 318, the conversion product and unreacted feed,if any, are separated from the catalyst by the separator 360. Desirably,the conversion product and unreacted feed, if any, are separated by aseries of cyclonic separators 362 as shown in FIG. 4. Further, at leasta portion of the catalyst exiting the riser reactors 330 contacts theimpingement device 380 and is deflected away from the second ends 334 ofthe riser reactors 330 to the quiescent zone 324.

Once the catalyst has been separated from the conversion product and theunreacted feed, if any, are removed from the shell 310 through the lines364 for further processing such as separation and purification. Aportion of the catalyst falls to the quiescent zone 324 in which thecatalyst is retained until it is removed from the shell 310. Thecatalyst is removed from the quiescent zone 324 through outlet 370 andcan be sent for regeneration in the catalyst regeneration apparatus 90.The function of the catalyst regeneration apparatus 90 is discussedabove in conjunction with the description of FIG. 1 and will not bediscussed in further detail here. A portion of the catalyst in thequiescent zone 324 will fall out of the quiescent zone 324 into thecatalyst return 338 and be returned to contact the feed.

Returning to FIG. 4, the remaining portion of the catalyst, after beingseparated from the products and unreacted feed, falls from the first end314 of shell 310 through the catalyst return 338 to the catalystretention zone 336. From the catalyst retention zone 336, the catalystis recycled for use in the hydrocarbon conversion reaction.

Representative embodiments of possible configurations of riser reactorsand catalyst returns are shown in cross section in FIG. 5. FIG. 5A showsa possible configuration for the riser reactors 20 for the hydrocarbonconversion apparatus 10 shown in FIG. 1. As shown in FIG. 5A, the riserreactors 20 are contained within a shell 26. If contained within a shell26, the area between the riser reactors and the shell 26 is filled withrefractory material 28. Useful refractory materials 28 include sand,cement, ceramic materials, high alumina bricks containing mullite orcorundum, high silica bricks, magnesite bricks, insulating firebrick ofclay or kaolin or any other high temperature resistant material.

FIG. 5B shows a cross section of a hydrocarbon conversion apparatussimilar to apparatus 10 shown in FIG. 1. In this embodiment, the riserreactors 20 are again contained within a shell 26. The shell 26 isfilled with refractory material 28 as described above. In thisembodiment, the catalyst returns are also contained within the shell 26and surrounded by the refractory material 28.

FIG. 5C shows a possible configuration for the riser reactors 220 shownin FIG. 3. In the embodiment shown in FIG. 5C, the catalyst return 250is shown as being centrally positioned in relation to the riser reactors220. The riser reactors 220 and the catalyst return 250 are containedwithin a shell 226. The area between the riser reactors and the shell226 is filled with refractory material 228. Useful refractory materialsare described above in conjunction with the description of FIG. 5A.

FIG. 5D shows a possible configuration for the riser reactors 330 shownin FIG. 4. As shown in FIG. 5D, the riser reactors 330 are centrallylocated within the shell 310. As described above in conjunction with thedescription of FIG. 4, the walls 331 of the riser reactors 330 and theshell 310 define the catalyst return 338. The area between the riserreactors 330 can optionally be filled with a first refractory material382. The shell 310 may also be optionally filled with a secondrefractory material 384. Useful refractory materials are described abovein conjunction with the description of FIG. 5A. With continuingreference to FIG. 5D, a person of skill in the art will appreciate thatthe first refractory material 382 and the second refractory material 384can be the same or different material.

FIG. 5E shows another possible configuration for the riser reactors 330shown in FIG. 4. As shown in FIG. 5E, the riser reactors 330 arecentrally located within the shell 310. In this embodiment, the riserreactors 330 are contained within a second shell 386 which has a wall388. The catalyst return 338 is defined by the wall 388 of the secondshell 386 and the shell 310. The areas between the walls 331 of theriser reactors 330 and the wall 388 of the second shell 386 are filledwith a first refractory material 390. The shell 310 may also be filledwith a second refractory material 392. Useful refractory materials aredescribed above in conjunction with the description of FIG. 5A. Withcontinuing reference to FIG. 5E, a person of skill in the art willappreciate that the first refractory material 390 and the secondrefractory material 392 can be the same or different material.

While the riser reactors and catalyst returns are shown in the variousFigures as having a circular cross section, the riser reactors andcatalyst returns may have any cross section which would facilitateoperation of the hydrocarbon conversion apparatus. Other useful crosssections for the riser reactors and the catalyst returns includeelliptical cross sections, polygonal cross sections and cross sectionsof sections of ellipses and polygons. Desirable cross-sections for theriser reactors and catalyst returns include circles and regular polygonswith sides of equal lengths. By “regular”, it is meant that the shape ofthe cross-section has no line segments with vertices, inside theboundaries of the shape, having angles greater than 180°. The mostdesirable cross-sections are circles, and triangles, squares, andhexagons with sides of equal length. The means of determiningcross-sectional areas for any cross-section shape is based on longestablished geometric principles well known to those skilled in the art.Similarly, desirable cross-sections for the separation zone includecircles and regular polygons with sides of equal lengths. The mostdesirable cross-sections are circles, and triangles, squares, andhexagons with sides of equal length.

While the position of the riser reactors relative to the separation zoneare shown in the figures as equidistant and symmetrical, alternateconfigurations are within the scope of the present invention. Forexample, the riser reactors may be positioned on one side of theseparation zone in a hemispherical layout. As another example, when theseparation zone has a circular or approximately circular cross-section,the riser reactors may be positioned in a line along the diameter theseparation zone. One skilled in the art will appreciate that a widevariety of configurations of the risers relative to the separation zonemay be utilized in the present invention.

One skilled in the art will further appreciate that the multiple riserreactors of the hydrocarbon conversion apparatus of the presentinvention may be formed by dividing a single riser reactor into aplurality of smaller riser reactors. For example, a larger, reactorhaving a circular cross section could be divided into several pie-shapedriser reactors. As another example, a riser reactor having a squarecross section could be divided into a plurality of riser reactors havingeither rectangular or smaller square cross sections.

The hydrocarbon conversion apparatus of the present invention is usefulto conduct most any hydrocarbon conversion process in which a fluidizedcatalyst is employed. Typical reactions include, for example, olefininterconversion reactions, oxygenate to olefin conversion reactions,oxygenate to gasoline conversion reactions, malaeic anhydrideformulation, vapor phase methanol synthesis, phthalic anhydrideformulation, Fischer Tropsch reactions, and acrylonitrile formulation.

The hydrocarbon conversion apparatus of the present invention isparticularly suited for conducting an oxygenate to olefin conversionreaction. In an oxygenate to olefin conversion reaction, an oxygenate isconverted to an olefin by contacting an oxygenate feed with a catalystunder conditions sufficient to convert the oxygenate to an olefin.

The process for converting oxygenates to light olefins employs a feedincluding an oxygenate. As used herein, the term “oxygenate” is definedto include, but is not necessarily limited to, hydrocarbons containingoxygen such as the following: aliphatic alcohols, ethers, carbonylcompounds (aldehydes, ketones, carboxylic acids, carbonates, and thelike), and mixtures thereof. The aliphatic moiety desirably shouldcontain in the range of from about 1–10 carbon atoms and more desirablyin the range of from about 1–4 carbon atoms. Representative oxygenatesinclude, but are not necessarily limited to, lower molecular weightstraight chain or branched aliphatic alcohols, and their unsaturatedcounterparts. Examples of suitable oxygenates include, but are notnecessarily limited to the following: methanol; ethanol; n-propanol;isopropanol; C₄–C₁₀ alcohols; methyl ethyl ether; dimethyl ether;diethyl ether; di-isopropyl ether; methyl formate; formaldehyde;di-methyl carbonate; methyl ethyl carbonate; acetone; and mixturesthereof. Desirably, the oxygenate used in the conversion reaction isselected from the group consisting of methanol, dimethyl ether andmixtures thereof. More desirably the oxygenate is methanol. The totalcharge of feed to the riser reactors may contain additional components,such as diluents.

One or more diluents may be fed to the riser reactors with theoxygenates, such that the total feed mixture comprises diluent in arange of from about 1 mol % and about 99 mol %. Diluents which may beemployed in the process include, but are not necessarily limited to,helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen,water, paraffins, other hydrocarbons (such as methane), aromaticcompounds, and mixtures thereof. Desired diluents include, but are notnecessarily limited to, water and nitrogen.

A portion of the feed may be provided to the reactor in liquid form.When a portion of the feed is provided in a liquid form, the liquidportion of the feed may be either oxygenate, diluent or a mixture ofboth. The liquid portion of the feed may be directly injected into theindividual riser reactors, or entrained or otherwise carried into theriser reactors with the vapor portion of the feed or a suitable carriergas/diluent. By providing a portion of the feed (oxygenate and/ordiluent) in the liquid phase, the temperature in the riser reactors canbe controlled. The exothermic heat of reaction of oxygenate conversionis partially absorbed by the endothermic heat of vaporization of theliquid portion of the feed. Controlling the proportion of liquid feed tovapor feed fed to the reactor is one possible method for controlling thetemperature in the reactor and in particular in the riser reactors.

The amount of feed provided in a liquid form, whether fed separately orjointly with the vapor feed, is from about 0.1 wt. % to about 85 wt. %of the total oxygenate content plus diluent in the feed. More desirably,the range is from about 1 wt. % to about 75 wt. % of the total oxygenateplus diluent feed, and most desirably the range is from about 5 wt. % toabout 65 wt. %. The liquid and vapor portions of the feed may be thesame composition, or may contain varying proportions of the same ordifferent oxygenates and same or different diluents. One particularlyeffective liquid diluent is water, due to its relatively high heat ofvaporization, which allows for a high impact on the reactor temperaturedifferential with a relatively small rate. Other useful diluents aredescribed above. Proper selection of the temperature and pressure of anyappropriate oxygenate and/or diluent being fed to the reactor willensure at least a portion is in the liquid phase as it enters thereactor and/or comes into contact with the catalyst or a vapor portionof the feed and/or diluent.

Optionally, the liquid fraction of the feed may be split into portionsand introduced to riser reactors a multiplicity of locations along thelength of the riser reactors. This may be done with either the oxygenatefeed, the diluent or both. Typically, this is done with the diluentportion of the feed. Another option is to provide a nozzle whichintroduces the total liquid fraction of the feed to the riser reactorsin a manner such that the nozzle forms liquid droplets of an appropriatesize distribution which, when entrained with the gas and solidsintroduced to the riser reactors, vaporize gradually along the length ofthe riser reactors. Either of these arrangements or a combinationthereof may be used to better control the temperature differential inthe riser reactors. The means of introducing a multiplicity of liquidfeed points in a reactor or designing a liquid feed nozzle to controldroplet size distribution is well known in the art and is not discussedhere.

The catalyst suitable for catalyzing an oxygenate-to-olefin conversionreaction includes a molecular sieve and mixtures of molecular sieves.Molecular sieves can be zeolitic (zeolites) or non-zeolitic(non-zeolites). Useful catalysts may also be formed from mixtures ofzeolitic and non-zeolitic molecular sieves. Desirably, the catalystincludes a non-zeolitic molecular sieve. Desired molecular sieves foruse with an oxygenate to olefins conversion reaction include “small” and“medium” pore molecular sieves. “Small pore” molecular sieves aredefined as molecular sieves with pores having a diameter of less thanabout 5.0 Angstroms. “Medium pore” molecular sieves are defined asmolecular sieves with pores having a diameter from about 5.0 to about10.0 Angstroms.

Useful zeolitic molecular sieves include, but are not limited to,mordenite, chabazite, erionite, ZSM-5, ZSM-34, ZSM-48 and mixturesthereof. Methods of making these molecular sieves are known in the artand need not be discussed here. Structural types of small pore molecularsieves that are suitable for use in this invention include AEI, AFT,APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO,KFI, LEV, LOV, LTA, MON, PAU, PHI, RHO, ROG, THO, and substituted formsthereof. Structural types of medium pore molecular sieves that aresuitable for use in this invention include MFI, MEL, MTW, EUO, MTT, HEU,FER, AFO, AEL, TON, and substituted forms thereof.

Silicoaluminophosphates (“SAPOs”) are one group of non-zeoliticmolecular sieves that are useful in an oxygenate to olefins conversionreaction. SAPOs comprise a three-dimensional microporous crystalframework structure of [SiO₂], [AlO₂] and [PO₂] tetrahedral units. Theway Si is incorporated into the structure can be determined by ²⁹Si MASNMR. See Blackwell and Patton, J. Phys. Chem., 92, 3965 (1988). Thedesired SAPO molecular sieves will exhibit one or more peaks in the ²⁹SiMAS NMR, with a chemical shift [(Si)] in the range of −88 to −96 ppm andwith a combined peak area in that range of at least 20% of the totalpeak area of all peaks with a chemical shift [(Si)] in the range of −88ppm to −115 ppm, where the [(Si)] chemical shifts refer to externaltetramethylsilane (TMS).

It is desired that the silicoaluminophosphate molecular sieve used insuch a process have a relatively low Si/Al₂ ratio. In general, the lowerthe Si/Al₂ ratio, the lower the C₁–C₄ saturates selectivity,particularly propane selectivity. A Si/Al₂ ratio of less than 0.65 isdesirable, with a Si/Al₂ ratio of not greater than 0.40 being preferred,and a SiAl₂ ratio of not greater than 0.32 being particularly preferred.

Silicoaluminophosphate molecular sieves are generally classified asbeing microporous materials having 8, 10, or 12 membered ringstructures. These ring structures can have an average pore size rangingfrom about 3.5–15 angstroms. Preferred are the small pore SAPO molecularsieves having an average pore size ranging from about 3.5 to 5angstroms, more preferably from 4.0 to 5.0 angstroms. These pore sizesare typical of molecular sieves having 8 membered rings.

In general, silicoaluminophosphate molecular sieves comprise a molecularframework of corner-sharing [SiO₂], [AlO₂], and [PO₂] tetrahedral units.This type of framework is effective in converting various oxygenatesinto olefin products.

Suitable silicoaluminophosphate molecular sieves for use in an oxygenateto olefin conversion process include SAPO-5, SAPO-8, SAPO-11, SAPO-16,SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36, SAPO-37,SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, the metalcontaining forms thereof, and mixtures thereof. Preferred are SAPO-18,SAPO-34, SAPO-35, SAPO-44, and SAPO-47, particularly SAPO-18 andSAPO-34, including the metal containing forms thereof, and mixturesthereof. As used herein, the term mixture is synonymous with combinationand is considered a composition of matter having two or more componentsin varying proportions, regardless of their physical state.

Additional olefin-forming molecular sieve materials can be mixed withthe silicoaluminophosphate catalyst if desired. Several types ofmolecular sieves exist, each of which exhibit different properties.Structural types of small pore molecular sieves that are suitable foruse in this invention include AEI, AFT, APC, ATN, ATT, ATV, AWW, BIK,CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO, KFI, LEV, LOV, LTA, MON, PAU,PHI, RHO, ROG, THO, and substituted forms thereof. Structural types ofmedium pore molecular sieves that are suitable for use in this inventioninclude MFI, MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, andsubstituted forms thereof. Preferred molecular sieves which can becombined with a silicoaluminophosphate catalyst include ZSM-5, ZSM-34,erionite, and chabazite.

Substituted SAPOs form a class of molecular sieves known as “MeAPSOs,”which are also useful in the present invention. Processes for makingMeAPSOs are known in the art. SAPOs with substituents, such as MeAPSOs,also may be suitable for use in the present invention. Suitablesubstituents, “Me,” include, but are not necessarily limited to, nickel,cobalt, manganese, zinc, titanium, strontium, magnesium, barium, andcalcium. The substituents may be incorporated during synthesis of theMeAPSOs. Alternately, the substituents may be incorporated aftersynthesis of SAPOs or MeAPSOs using many methods. These methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof.

Desired MeAPSOs are small pore MeAPSOs having pore size smaller thanabout 5 Angstroms. Small pore MeAPSOs include, but are not necessarilylimited to, NiSAPO-34, CoSAPO-34, NiSAPO-17, CoSAPO-17, and mixturesthereof.

Aluminophosphates (ALPOs) with substituents, also known as “MeAPOs,” areanother group of molecular sieves that may be suitable for use in anoxygenate to olefin conversion reaction, with desired MeAPOs being smallpore MeAPOs. Processes for making MeAPOs are known in the art. Suitablesubstituents include, but are not necessarily limited to, nickel,cobalt, manganese, zinc, titanium, strontium, magnesium, barium, andcalcium. The substituents may be incorporated during synthesis of theMeAPOs. Alternately, the substituents may be incorporated aftersynthesis of ALPOs or MeAPOs using many methods. The methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof.

The molecular sieve may also be incorporated into a solid composition,preferably solid particles, in which the molecular sieve is present inan amount effective to catalyze the desired conversion reaction. Thesolid particles may include a catalytically effective amount of themolecular sieve and matrix material, preferably at least one of a fillermaterial and a binder material, to provide a desired property orproperties, e.g., desired catalyst dilution, mechanical strength and thelike, to the solid composition. Such matrix materials are often to someextent porous in nature and often have some nonselective catalyticactivity to promote the formation of undesired products and may or maynot be effective to promote the desired chemical conversion. Suchmatrix, e.g., filler and binder, materials include, for example,synthetic and naturally occurring substances, metal oxides, clays,silicas, aluminas, silica-aluminas, silica-magnesias, silica-zirconias,silica-thorias, silica-beryllias, silica-titanias,silica-alumina-thorias, silica-aluminazirconias, and mixtures of thesematerials.

The solid catalyst composition preferably comprises about 1% to about99%, more preferably about 5% to about 90%, and still more preferablyabout 10% to about 80%, by weight of molecular sieve; and an amount ofabout 1% to about 99%, more preferably about 5% to about 90%, and stillmore preferably about 10% to about 80%, by weight of matrix material.

The preparation of solid catalyst compositions, e.g., solid particles,comprising the molecular sieve and matrix material, is conventional andwell known in the art and, therefore, is not discussed in detail here.

The catalyst may further contain binders, fillers, or other material toprovide better catalytic performance, attrition resistance,regenerability, and other desired properties. Desirably, the catalyst isfluidizable under the reaction conditions. The catalyst should haveparticle sizes of from about 20μ to about 3,000μ, desirably from about30μ to about 200μ, and more desirably from about 501μ to about 150μ. Thecatalyst may be subjected to a variety of treatments to achieve thedesired physical and chemical characteristics. Such treatments include,but are not necessarily limited to, calcination, ball milling, milling,grinding, spray drying, hydrothermal treatment, acid treatment, basetreatment, and combinations thereof.

Desirably, in an oxygenate to olefin conversion reaction conducted inthe hydrocarbon conversion apparatus of the present invention employs agas superficial velocity in the riser reactors of greater than 1 meterper second (m/s). As used herein and in the claims, the term, “gassuperficial velocity,” is defined as the volumetric flow rate ofvaporized feedstock, and any diluent, divided by the reactorcross-sectional area. Because the oxygenate is converted to a productincluding a light olefin while flowing through the reactor, the gassuperficial velocity may vary at different locations within the reactordepending on the total number of moles of gas present and the crosssection of a particular location in the reactor, temperature, pressure,and other relevant reaction parameters. The gas superficial velocity,including any diluents present in the feedstock, is maintained at a rategreater than 1 meter per second (m/s) at any point in the reactor.Desirably, the gas superficial velocity is greater than about 2 m/s.More desirably, the gas superficial velocity is greater than about 2.5m/s. Even more desirably, the gas superficial velocity is greater thanabout 4 m/s. Most desirably, the gas superficial velocity is greaterthan about 8 m/s.

Maintaining the gas superficial velocity at these rates increases theapproach to plug flow behavior of the gases flowing in the riserreactors. As the gas superficial velocity increases above 1 m/s, areduction in axial diffusion or back mixing of the gases results from areduction in internal recirculation of solids, which carry gas withthem. (Ideal plug flow behavior occurs when elements of the homogeneousfluid reactant move through a reactor as plugs moving parallel to thereactor axis). Minimizing the back mixing of the gases in the reactorincreases the selectivity to the desired light olefins in the oxygenateconversion reaction.

When the gas superficial velocity approaches 1 m/s or higher, asubstantial portion of the catalyst in the reactor may be entrained withthe gas exiting the riser reactors. At least a portion of the catalystexiting the riser reactors is recirculated to recontact the feed throughthe catalyst return.

Desirably, the rate of catalyst, comprising molecular sieve and anyother materials such as binders, fillers, etc., recirculated torecontact the feed is from about 1 to about 100 times, more desirablyfrom about 10 to about 80 times, and most desirably from about 10 toabout 50 times the total feed rate, by weight, of oxygenates to thereactor.

The temperature useful to convert oxygenates to light olefins variesover a wide range depending, at least in part, on the catalyst, thefraction of regenerated catalyst in a catalyst mixture, and theconfiguration of the reactor apparatus and the reactor. Although theseprocesses are not limited to a particular temperature, best results areobtained if the process is conducted at a temperature from about 200° C.to about 700° C., desirably from about 250° C. to about 600° C., andmost desirably from about 300° C. to about 500° C. Lower temperaturesgenerally result in lower rates of reaction, and the formation rate ofthe desired light olefin products may become markedly slower. However,at temperatures greater than 700° C., the process may not form anoptimum amount of light olefin products, and the rate at which coke andlight saturates form on the catalyst may become too high.

Light olefins will form—although not necessarily in optimum amounts—at awide range of pressures including, but not limited to, pressures fromabout 0.1 kPa to about 5 MPa. A desired pressure is from about 5 kPa toabout 1 MPa and most desirably from about 20 kPa to about 500 kPa. Theforegoing pressures do not include that of a diluent, if any, and referto the partial pressure of the feed as it relates to oxygenate compoundsand/or mixtures thereof. Pressures outside of the stated ranges may beused and are not excluded from the scope of the invention. Lower andupper extremes of pressure may adversely affect selectivity, conversion,coking rate, and/or reaction rate; however, light olefins will stillform and, for that reason, these extremes of pressure are consideredpart of the present invention.

A wide range of WHSVs for the oxygenate conversion reaction, defined asweight of total oxygenate fed to the riser reactors per hour per weightof molecular sieve in the catalyst in the riser reactors, function withthe present invention. The total oxygenate fed to the riser reactorsincludes all oxygenate in both the vapor and liquid phase. Although thecatalyst may contain other materials which act as inerts, fillers orbinders, the WHSV is calculated using only the weight of molecular sievein the catalyst in the riser reactors. The WHSV is desirably high enoughto maintain the catalyst in a fluidized state under the reactionconditions and within the reactor configuration and design. Generally,the WHSV is from about 1 hr⁻¹ to about 5000 hr⁻¹, desirably from about 2hr⁻¹ to about 3000 hr⁻¹, and most desirably from about 5 hr⁻¹ to about1500 hr⁻¹. The applicants have discovered that operation of theoxygenate to olefin conversion reaction at a WHSV greater than 20 hr⁻¹reduces the methane content in the product slate of the conversionreaction. Thus, the conversion reaction is desirably operated at a WHSVof at least about 20 hr⁻¹. For a feed comprising methanol, dimethylether, or mixtures thereof, the WHSV is desirably at least about 20 hr⁻¹and more desirably from about 20 hr⁻¹ to about 300 hr^(−1.)

It is particularly preferred that the reaction conditions for makingolefins from an oxygenate comprise a WHSV of at least about 20 hr⁻¹ anda Temperature Corrected Normalized Methane Selectivity (TCNMS) of lessthan about 0.016. As used herein, TCNMS is defined as the NormalizedMethane Selectivity (NMS) when the temperature is less than 400° C. TheNMS is defined as the methane product yield divided by the ethyleneproduct yield wherein each yield is measured on or is converted to aweight % basis. When the temperature is 400° C. or greater, the TCNMS isdefined by the following equation, in which T is the average temperaturewithin the reactor in ° C:

${TCNMS} = {\frac{NMS}{1 + \left( {\left( {\left( {T - 400} \right)/400} \right) \times 14.84} \right)}.}$

Oxygenate conversion should be maintained sufficiently high to avoid theneed for commercially unacceptable levels of feed recycling. While 100%oxygenate conversion is desired for the purpose of completely avoidingfeed recycle, a reduction in unwanted by-products is observed to as muchas about 50% of the feed can be commercially acceptable, conversionrates from about 50% to about 98% are desired. Conversion rates may bemaintained in this range—50% to about 98%—using a number of methodsfamiliar to persons of ordinary skill in the art. Examples include, butare not necessarily limited to, adjusting one or more of the following:reaction temperature; pressure; flow rate (weight hourly space velocityand/or gas superficial velocity); catalyst recirculation rate; reactorapparatus configuration; reactor configuration; feed composition; amountof liquid feed relative to vapor feed (as will be discussed below);amount of recirculated catalyst; degree of catalyst regeneration; andother parameters which affect the conversion.

During the conversion of oxygenates to light olefins, carbonaceousdeposits accumulate on the catalyst used to promote the conversionreaction. At some point, the build up of these carbonaceous depositscauses a reduction in the capability of the catalyst to convert theoxygenate feed to light olefins. At this point, the catalyst ispartially deactivated. When a catalyst can no longer convert anoxygenate to an olefin product, the catalyst is considered to be fullydeactivated. As an optional step in an oxygenate to olefin conversionreaction, a portion of the catalyst is withdrawn from the reactor and atleast a portion of the portion removed from the reactor is partially, ifnot fully, regenerated in a regeneration apparatus, such as regenerationapparatus 80 as shown in FIG. 4. By regeneration, it is meant that thecarbonaceous deposits are at least partially removed from the catalyst.Desirably, the portion of the catalyst withdrawn from the reactor is atleast partially deactivated. The remaining portion of the catalyst inthe reactor is re-circulated without regeneration, as described above.The regenerated catalyst, with or without cooling, is then returned tothe reactor. Desirably, the rate of withdrawing the portion of thecatalyst for regeneration is from about 0.1% to about 99% of the rate ofthe catalyst exiting the reactor. More desirably, the rate is from about0.2% to about 50%, and, most desirably, from about 0.5% to about 5%.

Desirably, a portion of the catalyst, comprising molecular sieve and anyother materials such as binders, fillers, etc., is removed from thereactor for regeneration and recirculation back to the reactor at a rateof from about 0.1 times to about 10 times, more desirably from about 0.2to about 5 times, and most desirably from about 0.3 to about 3 times thetotal feed rate of oxygenates to the reactor. These rates pertain to thecatalyst containing molecular sieve only, and do not includenon-reactive solids. The rate of total solids, i.e., catalyst andnon-reactive solids, removed from the reactor for regeneration andrecirculation back to the reactor will vary these rates in directproportion to the content of non-reactive solids in the total solids.

Desirably, the catalyst regeneration is carried out in a regenerationapparatus in the presence of a gas comprising oxygen or other oxidants.Examples of other oxidants include, but are not necessarily limited to,singlet O₂, O₃, SO₃, N₂O, NO, NO₂, N₂O₅, and mixtures thereof. Air andair diluted with nitrogen or CO₂ are desired regeneration gases. Theoxygen concentration in air can be reduced to a controlled level tominimize overheating of, or creating hot spots in, the regenerator. Thecatalyst may also be regenerated reductively with hydrogen, mixtures ofhydrogen and carbon monoxide, or other suitable reducing gases.

The catalyst may be regenerated in any number of methods—batch,continuous, semi-continuous, or a combination thereof. Continuouscatalyst regeneration is a desired method. Desirably, the catalyst isregenerated to a level of remaining coke from about 0.01 wt % to about15 wt % of the weight of the catalyst.

The catalyst regeneration temperature should be from about 250° C. toabout 750° C., and desirably from about 500° C. to about 700° C. Becausethe regeneration reaction takes place at a temperature considerablyhigher than the oxygenate conversion reaction, it may be desirable tocool at least a portion of the regenerated catalyst to a lowertemperature before it is sent back to the reactor. A heat exchanger, notshown, located external to the regeneration apparatus may be used toremove some heat from the catalyst after it has been withdrawn from theregeneration apparatus. When the regenerated catalyst is cooled, it isdesirable to cool it to a temperature which is from about 200° C. higherto about 200° C. lower than the temperature of the catalyst withdrawnfrom the reactor. More desirably, the regenerated catalyst is cooled toa temperature from about 10° C. to about 200° C. lower than thetemperature of the catalyst withdrawn from the reactor. This cooledcatalyst then may be returned to either some portion of the reactor, theregeneration apparatus, or both. When the regenerated catalyst from theregeneration apparatus is returned to the reactor, it may be returned toany portion of the reactor. It may be returned to the catalystcontainment area to await contact with the feed, the separation zone tocontact products of the feed or a combination of both.

Desirably, catalyst regeneration is carried out at least partiallydeactivated catalyst that has been stripped of most of readily removableorganic materials (organics) in a stripper or stripping chamber first.This stripping can be achieved by passing a stripping gas over the spentcatalyst at an elevated temperature. Gases suitable for strippinginclude steam, nitrogen, helium, argon, methane, CO₂, CO, hydrogen, andmixtures thereof. A preferred gas is steam. Gas hourly space velocity(GHSV, based on volume of gas to volume of catalyst and coke) of thestripping gas is from about 0.1 h⁻¹ to about 20,000 h⁻¹. Acceptabletemperatures of stripping are from about 250° C. to about 750° C., anddesirably from about 350° C. to about 675° C.

The method of making the preferred olefin product in this invention caninclude the additional step of making the oxygenate compositions fromhydrocarbons such as oil, coal, tar sand, shale, biomass and naturalgas. Methods for making the compositions are known in the art. Thesemethods include fermentation to alcohol or ether, making synthesis gas,then converting the synthesis gas to alcohol or ether. Synthesis gas canbe produced by known processes such as steam reforming, autothermalreforming and partial oxidization.

One skilled in the art will also appreciate that the olefins produced bythe oxygenate-to-olefin conversion reaction of the present invention canbe polymerized to form polyolefins, particularly polyethylene andpolypropylene. Processes for forming polyolefins from olefins are knownin the art. Catalytic processes are preferred. Particularly preferredare metallocene, Ziegler/Natta and acid catalytic systems. See, forexample, U.S. Pat. Nos. 3,258,455; 3,305,538; 3,364,190; 5,892,079;4,659,685; 4,076,698; 3,645,992; 4,302,565; and 4,243,691, the catalystand process descriptions of each being expressly incorporated herein byreference. In general, these methods involve contacting the olefinproduct with a polyolefin-forming catalyst at a pressure and temperatureeffective to form the polyolefin product.

A preferred polyolefin-forming catalyst is a metallocene catalyst. Thepreferred temperature range of operation is between 50° C. and 240° C.and the reaction can be carried out at low, medium or high pressure,being anywhere from 1 bar to 200 bars. For processes carried out insolution, an inert diluent can be used, and the preferred operatingpressure range is between 10 and 150 bars, with a preferred temperaturebetween 120° C. and 230° C. For gas phase processes, it is preferredthat the temperature generally be from 60° C. to 160° C., and that theoperating pressure be from 5 bars to 50 bars.

In addition to polyolefins, numerous other olefin derivatives may beformed from the olefins produced by the process of the present inventionor olefins recovered therefrom. These include, but are not limited to,aldehydes, alcohols, acetic acid, linear alpha olefins, vinyl acetate,ethylene dichloride and vinyl chloride, ethylbenzene, ethylene oxide,ethylene glycol, cumene, isopropyl alcohol, acrolein, allyl chloride,propylene oxide, acrylic acid, ethylene-propylene rubbers, andacrylonitrile, and trimers and dimers of ethylene, propylene orbutylenes. The methods of manufacturing these derivatives are well knownin the art, and therefore are not discussed here.

Persons of ordinary skill in the art will recognize that manymodifications may be made to the present invention without departingfrom the spirit and scope of the present invention. The embodimentsdescribed herein are meant to be illustrative only and should not betaken as limiting the invention, which is defined by the followingclaims.

1. An oxygenate to olefin conversion process comprising: (a) feeding anoxygenate feed to a plurality of necks, each neck having a headextending into an associated riser reactor, wherein each riser reactorextends into a shell and into a separation zone of a single hydrocarbonconversion apparatus; (b) reacting said feed by contacting with anon-zeolitic molecular sieve catalyst from a catalyst retention zone ina mouth of each riser reactor at a gas superficial velocity of greaterthan 1 meter per second at any point in the conversion apparatus,wherein each head extending into each riser is positioned at or abovethe mouth, and the reaction of said feed produces a light olefinproduct; (c) separating said catalyst from said product in theseparation zone; (d) transporting a portion of said catalyst from saidshell of the separation zone to a regenerator apparatus; (e) moving aportion of said catalyst from said shell of the separation zone to thecatalyst retention zone by way of a catalyst return that defines atleast a portion of the catalyst retention zone; and (f) returningcatalyst from the regenerator apparatus to the separation zone, thecatalyst retention zone, the catalyst return or any combination thereof.2. The process of claim 1 wherein said feed is fed to each of saidplurality of riser reactors in a substantially equal amount.
 3. Theprocess of claim 2 wherein said feed is fed to each of said plurality ofriser reactors such that the flow of feed to each reactor varies by nomore than 25%, by volume rate, from one of said plurality of riserreactors to another of said plurality of riser reactors.
 4. The processof claim 3 wherein said feed is fed to each of said plurality of riserreactors such that the flow of feed to each reactor varies by no morethan 10%, by volume rate, from one of said plurality of riser reactorsto another of said plurality of riser reactors.
 5. The process of claim4 wherein said feed is fed to each of said plurality of riser reactorssuch that the flow of feed to each reactor varies by no more than 1%, byvolume rate, from one of said plurality of riser reactors to another ofsaid plurality of riser reactors.
 6. The process of claim 2 wherein saidfeed is fed to each of said plurality of riser reactors such that theflow of feed to each reactor varies by no more than 25%, by mass percentfor each component in the feed, from one of said plurality of riserreactors to another of said plurality of riser reactors.
 7. The processof claim 6 wherein said feed is fed to each of said plurality of riserreactors such that the flow of feed to each reactor varies by no morethan 10%, by mass percent for each component in the feed, from one ofsaid plurality of riser reactors to another of said plurality of riserreactors.
 8. The process of claim 7 wherein said feed is fed to each ofsaid plurality of riser reactors such that the flow of feed to eachreactor varies by no more than 1%, by mass percent for each component inthe feed, from one of said plurality of riser reactors to another ofsaid plurality of riser reactors.
 9. The process of claim 1 wherein saidcatalyst is separated from said product with a separator selected fromthe group consisting of cyclonic separators, filters, screens,impingement devices, plates, cones and combinations thereof.
 10. Theprocess of claim 1 further including the step of stripping said portionof said catalyst transported from the separation zone to the regeneratorapparatus prior to entering the regenerator apparatus.
 11. The processof claim 1 wherein each of said riser reactors has a height of from 10meters to 70 meters.
 12. The process of claim 11 wherein each of saidriser reactors has a width of from one meter to three meters.
 13. Theprocess of claim 1 wherein each of said riser reactors has a crosssectional area of no greater than 12 m².
 14. The process of claim 13wherein each of said riser reactors has a cross sectional area of nogreater than 7 m².
 15. The process of claim 14 wherein each of saidriser reactors has a cross sectional area or no greater than 3.5 m². 16.The process of claim 1 wherein each of said riser reactors has a crosssectional area and said cross sectional area of one of said riserreactors varies by no more than 20% from the cross sectional area ofanother of said riser reactors.
 17. The process of claim 16 wherein saidcross sectional area of one of said riser reactors varies by no morethan 10% from the cross sectional area of another of said riserreactors.
 18. The process of claim 17 wherein said cross sectional areaof one of said riser reactors varies by no more than 1% from the crosssectional area of another of said riser reactors.
 19. The process ofclaim 1 wherein said catalyst is a silicoaluminophosphate molecularsieve.
 20. The process of claim 1 wherein said feed is selected from thegroup of methanol; ethanol; n-propanol; isopropanol; C₄–C₁₀ alcohols;methyl ethyl ether; dimethyl ether; diethyl ether; di-isopropyl ether;methyl formate; formaldehyde; di-methyl carbonate; methyl ethylcarbonate; acetone; and mixtures thereof.